Method for removing oxygenates, olefins, heavy hydrocarbons and waste water from a reactor effluent using a quench tower

ABSTRACT

The present invention relates to a process for catalytically converting a feedstock comprising an oxygenate to olefins utilizing a heat exchange device to transfer heat from at least a portion of an effluent of an oxygenate conversion reactor to the feedstock to cause at least a portion of the feedstock to vaporize.

RELATED APPLICATIONS

[0001] This application is a continuation-in-part of co-pending U.S.application Ser. No. 09/069,403, filed Apr. 29, 1998, now allowed.

FIELD OF THE INVENTION

[0002] The present invention relates to a process for increasing theefficiency of heat recovery and improving heat integration with directproduct quenching in the selective conversion of oxygenates to olefins.

BACKGROUND OF THE INVENTION

[0003] Light olefins (defined herein as ethylene, propylene, butenes andmixtures thereof) serve as feeds for the production of numerousimportant chemicals and polymers. Light olefins traditionally areproduced by cracking petroleum feeds. Because of the limited supply andescalating cost of petroleum feeds, the cost of producing olefins frompetroleum sources has increased steadily. Efforts to develop and improveolefin production technologies, particularly light olefins productiontechnologies, based on alternative feedstocks have increased.

[0004] An important type of alternative feedstocks for the production oflight olefins are oxygenates, such as alcohols, particularly methanoland ethanol, dimethyl ether, methyl ethyl ether, methyl formate, anddimethyl carbonate. Alcohols may be produced by fermentation, or fromsynthesis gas derived from natural gas, petroleum liquids, carbonaceousmaterials including coal, recycled plastics, municipal wastes,agricultural products, or most organic materials. Because of the widevariety of raw material sources, alcohol, alcohol derivatives, and otheroxygenates have promise as an economical, non-petroleum feedstock sourcefor olefin production.

[0005] The conversion of oxygenates to olefins takes place at arelatively high temperature, generally higher than about 250° C.,preferably higher than about 300° C. Because the conversion reaction isexothermic, the effluent typically has a higher temperature than theinitial temperature in the reactor. Many methods and/or process schemeshave been proposed to manage the heat of reaction generated from theoxygenate conversion reaction inside of the reactor in order to avoidtemperature surges and hot spots, and thereby to reduce the rate ofcatalyst deactivation and reduce the production of undesirable products,such as methane, ethane, carbon monoxide and carbonaceous deposits orcoke. It would be very useful to have a process that effectivelyutilizes the heat of reaction contained in the products exiting theoxygenate conversion reactor, optimizes heat recovery, and reducesoverall utility consumption in the conversion of oxygenates to olefins.Such a process is environmentally, economically, and commercially moreattractive.

[0006] In the conventional systems, the oxygenate conversion reaction ispredominantly conducted in the vapor phase using feedstocks and diluentsthat are usually liquid at ambient conditions. This requires supplyingsubstantial heat to the process to boil the oxygenate feedstock prior tointroducing it to the reactor, conventionally supplied by steam heatexchange or furnaces. Loss of energy is incurred in these indirect heatexchange methods, and substantial equipment is required. For steam,boilers must be built in addition to a steam/feed exchanger, andconstruction of a furnace is more expensive and complicated than atraditional heat exchanger. Methods are needed to improve the energyefficiency of the oxygenate conversion process and reduce the cost ofproviding vaporized oxygenate feedstock to an oxygenate conversionreactor.

[0007] Energy efficiency and cost of providing vaporized oxygenate feedis further complicated if utilization of a diluent is desired. The mostcommonly noted diluent, water/steam as disclosed in U.S. Pat. No.5,714,662, requires substantial energy and equipment cost to generate,but has the advantage of being easily able to separate from desiredlight olefins (especially ethylene and propylene). Other very commonlynoted diluents such as inert gases, including nitrogen, helium and evenmethane (see U.S. Pat. No. 5,744,680) require no energy or equipment tovaporize, but require extensive energy and equipment to separate fromthe desired light olefin product. Further, use of diluents can allowhigh total pressures while providing low oxygenate partial pressures,which can be advantageous in reducing compression energy needed in theoverall (including olefin separation and recovery) oxygenate conversionprocess, but this benefit may be outweighed by the energy costs ofboiling and separating of the diluent just noted. Proper selection ofdiluent composition is also needed to improve the energy efficiency inthe overall process and reduce the cost of providing vaporized oxygenatefeedstock to an oxygenate conversion reactor.

SUMMARY OF THE INVENTION

[0008] The present invention provides a process for converting anoxygenate to olefins with increased heat recovery and heat integration,said process comprising: heating a feedstock comprising said oxygenatehaving a first heat content from a first temperature to a secondtemperature through from one to about three stages having successivelyhigher heat contents; contacting said feedstock at said secondtemperature with a catalyst comprising a molecular sieve underconditions effective to produce a deactivated catalyst havingcarbonaceous deposits and a product comprising said olefins, whereinsaid molecular sieve comprises pores having a pore diameter smaller thanabout 10 Angstroms and said product has a third temperature which ishigher than said second temperature; quenching said product with amedium at an initial temperature and in an amount sufficient for forminga light product fraction and a heavy product fraction wherein said lightproduct fraction comprises light olefins and said heavy product fractionhas a final temperature which is higher than said first temperature byat least about 5° C.; using said heavy product fraction to provide heatat one or more of said stages to achieve said higher heat contents.

[0009] In another preferred embodiment, the process for converting anoxygenate to olefins comprises providing a feedstock comprising theoxygenate, transferring heat from at least a portion of an effluent ofan oxygenate conversion reactor to the feedstock to cause at least aportion of the feedstock to vaporize and form a vaporized feedstock, andcontacting the vaporized feedstock at a temperature from about 200 toabout 750° C. and a pressure from 1 kPa to 100 MPa with a catalystcomprising a molecular sieve having a pore diameter smaller than 10Angstroms, wherein the feedstock has a boiling range of no greater thanabout 30° C., the oxygenate conversion reactor converts at least aportion of the feedstock into the effluent and the effluent comprisesthe olefins.

[0010] In another preferred embodiment, the process for converting anoxygenate to olefins comprises providing a feedstock comprising theoxygenate and a diluent, transferring heat from at least a portion of aneffluent of an oxygenate conversion reactor to the feedstock to cause atleast a portion of the feedstock to vaporize and form a vaporizedfeedstock, and separating the diluent from the effluent, wherein thefeedstock has a boiling range of no greater than about 30° C., theoxygenate conversion reactor converts at least a portion of thefeedstock into the effluent, and the effluent comprises the olefins.

[0011] In another preferred embodiment, the process for converting anoxygenate to olefins comprises providing a feedstock comprising theoxygenate and a diluent, transferring heat from at least a portion of aneffluent of an oxygenate conversion reactor to the feedstock to cause atleast a portion of the feedstock to vaporize and form a vaporizedfeedstock, contacting the vaporized feedstock at a temperature fromabout 200 to about 750° C. and a pressure from 1 kPa to 100 MPa with acatalyst comprising a molecular sieve having a pore diameter smallerthan 10 Angstroms, and separating the diluent from the effluent, whereinthe feedstock has a boiling range of no greater than about 30° C., theoxygenate conversion reactor converts at least a portion of thefeedstock into the effluent, and the effluent comprises the olefins.

BRIEF DESCRIPTION OF THE DRAWINGS

[0012]FIG. 1 is a flow diagram of a preferred embodiment of increasingheat recovery in the present invention.

DETAILED DESCRIPTION OF THE INVENTION

[0013] The present invention provides a process for increasing heatrecovery and decreasing energy and utility requirements during theconversion of oxygenates to olefins. The process involves taking theproduct mixture, including any unreacted oxygenate feed, from anoxygenate conversion reactor and, without fractionating the products,directly quenching the product mixture with a suitable medium,preferably water. This type of quenching hereinafter will be referred toas “direct product quench.” The direct product quench removes heat fromthe product mixture, causing higher boiling components, such as waterand unreacted oxygenate feed, to condense and form a heavy productfraction. The heavy product fraction is separated from the light productfraction comprising gaseous hydrocarbon components such as lightolefins, methane, ethane, propane, and butanes. The heavy productfraction may be divided further into several fractions. The heavyproduction fraction, or any, or all of the several fractions may besubjected to various techniques or methods to separate the quench mediumfrom other components. The heavy product fraction, or any, or all of theseveral fractions or streams produced from quench medium separationsthereof, may be used to supply at least part of the heat needed tovaporize or otherwise to increase the heat content of the oxygenatefeedstock, through from one to about three stages, prior to beingintroduced into the oxygenate conversion reactor for contacting theoxygenate conversion catalyst. These stages give the oxygenate feedstocksuccessively higher heat content.

[0014] Most catalysts that are used in oxygenate conversion processescomprise molecular sieves, both zeolitic (zeolites) and non-zeolitictypes. The present invention should achieve many of the desiredimprovements using substantially any molecular sieve catalyst,regardless of the structure type or pore size. Mixtures of zeolitic andnon-zeolitic molecular sieves also may be used. Preferred molecularsieve catalysts for use according to the present invention comprise“small” and “medium” pore molecular sieve catalysts. “Small pore”molecular sieve catalysts are defined as catalysts with pores having apore diameter of less than about 5.0 Angstroms. “Medium pore” molecularsieve catalysts are defined as catalysts with pores having a porediameter in the range of from about 5.0 to about 10.0 Angstroms. “Largepore” molecular sieve catalysts are catalysts with pores having a porediameter larger than about 10.0 Angstroms. Generally, large poremolecular sieve catalysts, without additional appropriate modificationsand/or treatments, are not preferred catalysts for converting oxygenatesto light olefins.

[0015] Zeolitic molecular sieve catalysts suitable for the use in thepresent invention with varying degree of effectiveness include, but arenot necessarily limited to AEI, AFI, CHA, ERI, LOV, RHO, THO, MFI, FER,and substituted examples of these structural types, as described in W.M. Meier and D. H. Olson, Atlas of Zeolitic Structural Types(Butterworth Heineman-third edition, 1997), incorporated herein byreference. Preferred zeolite catalysts include but are not necessarilylimited to zeolite 3 A, zeolite 4A, zeolite 5 A (collectively referredto hereinafter as zeolite A), ZK-5, ZSM-5, ZSM-34, erionite, chabazite,offretite, silicalite, borosilicates and mixtures thereof. See Meier andOlson. These zeolites may be obtained from many companies and commercialsources such as Mobil, AMOCO, UCI, Engelhard, Aldrich Chemical Company,Johnson Matthey Company, Union Carbide Corporation, and others.

[0016] Silicoaluminophosphates (“SAPO's”) are one group of non-zeoliticmolecular sieves that are useful in the present invention. SuitableSAPO's for use in the invention include, but are not necessarily limitedto SAPO-17, SAPO-18, SAPO-34, SAPO-44, and mixtures thereof Small poreSAPO's are preferred for producing light olefins. A preferred SAPO isSAPO-34, which may be synthesized according to U.S. Pat. No. 4,440,871.incorporated herein by reference, and Zeolites, Vol. 17, pp. 212-222(1996), incorporated herein by reference. SAPO-18 may be synthesizedaccording to J. Chen et al. Studies in Surface Sciences and Catalysis,Proceedings of the Tenth International Catalysis Society, Volume 84, pp17-31 (1994).

[0017] Substituted silicoaluminophosphates (substituted SAPO's) formanother class of non-zeolitic molecular sieves known as “MeAPSO's,” thatare suitable for use as catalysts in the present invention. MeAPSO's aredescribed in U.S. Pat. No. 4,567,029 and U.S. Pat. No. 5,126,308,incorporated herein by reference. SAPO's with substituents incorporatedafter synthesis, also may be suitable for use in the present invention.Suitable substituents, “Me,” include, but are not necessarily limited tonickel, cobalt, manganese, chromium, iron, zinc, strontium, magnesium,barium, and calcium. Preferred MeAPSO's include, but are not necessarilylimited to NiSAPO-17, NiSAPO-34, Co-SAPO-17, Co-SAPO-34, Sr modifiedSAPO-17 (SrAPSO-17), Sr modified SAPO-18 (SrAPSO-18), Sr modifiedSAPO-34 (SrAPSO-34), SrAPSO-44, and mixtures thereof Differentsubstituents may be incorporated during or after the synthesis of thesilicoaluminophosphates.

[0018] Substituted aluminophosphates (ALPO) known as MeAPO's may also beused as the non-zeolitic molecular sieve catalysts for the presentinvention. MeAPO's include, but are not necessarily limited to ZnAPO,ZrAPO, TiAPO, and mixtures thereof These molecular sieves may besynthesized according to U.S. Pat. Nos. 4,861,743, 4,567,029, 5,126,308.

[0019] Because the catalyst may be used in a variety of oxygenateconversion reactors and/or under a variety of reaction conditions, itmay contain binders, fillers, or other material to provide bettercatalytic performance, attrition resistance, regenerability, and otherdesired properties for a particular type reactor. When used in afluidized bed reactor, the catalyst should be fluidizable under thereaction conditions. A catalyst may be subjected further to a variety oftreatments to achieve the desired physical, mechanical, and catalyticcharacteristics. Such treatments include, but are not necessarilylimited to calcination, milling, ball milling, grinding, spray drying,hydrothermal treatment with steam at elevated temperatures-from about400° C. to about 800° C., acid treatment, base treatment, andcombinations thereof.

[0020] The oxygenate feedstock of this invention comprises at least oneorganic compound which contains at least one oxygen atom, such asaliphatic alcohols, ethers, carbonyl compounds (aldehydes, ketones,carboxylic acids, carbonates, esters and the like). When the oxygenateis an alcohol, the alcohol can include an aliphatic moiety having from 1to 10 carbon atoms, more preferably from 1 to 4 carbon atoms.

[0021] Representative alcohols include but are not necessarily limitedto lower straight and branched chain aliphatic alcohols and theirunsaturated counterparts. Examples of suitable oxygenate compoundsinclude, but are not limited to: methanol; ethanol; n-propanol;isopropanol; C₄-C₂₀ alcohols; methyl ethyl ether; dimethyl ether;diethyl ether; di-isopropyl ether; formaldehyde; dimethyl carbonate;dimethyl ketone; acetic acid; and mixtures thereof Preferred oxygenatecompounds are methanol, dimethyl ether, or a mixture thereof.

[0022] The method of making the preferred olefin product in thisinvention can include the additional step of making these compositionsfrom hydrocarbons such as oil, coal, tar sand, shale, biomass andnatural gas. Methods for making the compositions are known in the art.These methods include fermentation to alcohol or ether, making synthesisgas, then converting the synthesis gas to alcohol or ether. Synthesisgas can be produced by known processes such as steam reforming,autothermal reforming and partial oxidization.

[0023] One or more inert diluents may be present in the feedstock, forexample, in an amount of from 1 to 99 molar percent, based on the totalnumber of moles of all feed and diluent components fed to the reactionzone (or catalyst). As defined herein, diluents are compositions whichare essentially non-reactive across a molecular sieve catalyst, andprimarily function to make the oxygenates in the feedstock lessconcentrated. Typical diluents include, but are not necessarily limitedto helium, argon, nitrogen, carbon monoxide, carbon dioxide, water,essentially non-reactive paraffins (especially the alkanes such asmethane, ethane, and propane), essentially non-reactive alkylenes,essentially non-reactive aromatic compounds, and mixtures thereof Thepreferred diluents are water and nitrogen. Water can be injected ineither liquid or vapor form.

[0024] Hydrocarbons can also be included as part of the feedstock, i.e.,as co-feed. As defined herein, hydrocarbons included with the feedstockare hydrocarbon compositions which are converted to another chemicalarrangement when contacted with molecular sieve catalyst. Thesehydrocarbons can include olefins, reactive paraffins, reactivealkylaromatics, reactive aromatics or mixtures thereof. Preferredhydrocarbon co-feeds include, propylene, butylene, pentylene, C₄⁺hydrocarbon mixtures, C₅ ⁺hydrocarbon mixtures, and mixtures thereofMore preferred as co-feeds are a C₄ ⁺hydrocarbon mixtures, with the mostpreferred being C₄ ⁺hydrocarbon mixtures which are obtained fromseparation and recycle of reactor product.

[0025] Preferably, the oxygenate feedstock should be at least partiallyvaporized and contacted in a suitable oxygenate conversion reactor withthe selected molecular sieve catalyst under process conditions effectiveto produce the desired olefins at an acceptable conversion level withdesired selectivities.

[0026] The temperature employed in the conversion process may vary overa wide range depending, at least in part, on the pressure, the selectedcatalyst, the reactor configuration, the weight hourly space velocity,and other reaction parameters. Although not limited to a particulartemperature, best results will be obtained if the process is conductedat temperatures in the range of from about 200° C. to about 750° C.,preferably in the range of from about 250° C. to about 650° C., and mostpreferably in the range of from about 300° C. to about 600° C.

[0027] Since the oxygenate feedstock normally is stored at ambienttemperatures before it is used in the conversion process, the feedstockhas to be heated to a higher temperature with a much higher heat contentsuitable for contacting the oxygenate conversion catalyst. It ispreferable to increase the heat content and/or the temperature of thefeedstock through from one to about three intermediate stages, with eachstage having a successively higher heat content. Many different streamsin the oxygenate conversion process may be suitable sources forproviding the necessary heat to increase heat contents. These streams,including those derived from the heavy product fraction from the quenchtower and the streams from the fractionator separating quench mediumfrom other components, are described in more detail below. It should bepointed out that a stream may have a higher heat content after a heatexchange even though it has a lower temperature, primarily resultingfrom pressure changes and/or phase changes, such as vaporization of aliquid. The pressure and/or phase changes are needed for the oxygenateconversion process.

[0028] Light olefin products will form——although not necessarily inoptimum amounts——at a wide range of pressures, including but notnecessarily limited to sub- and super-atmospheric pressures andautogeneous pressures, ——in the range of from about 1 kPa to about 100MPa. A preferred pressure is in the range of from about 5 kPa to about50 MPa, most preferably in the range of from about 50 kPa to about 500kPa. The foregoing pressures are exclusive of diluent, if any ispresent, and refer to the partial pressure of the feedstock as itrelates to oxygenate compounds and/or mixtures thereof. Pressuresoutside of the stated ranges may be used and are not excluded from thescope of the invention.

[0029] A steady state or semi-steady state production of light olefinproducts may be attained and/or sustained over a period of time, largelydetermined by the reactor type, the reactor configuration, thetemperature, the pressure, the catalyst selected, the amount of spentcatalyst recirculated (if any), the level of catalyst regeneration, theamount of carbonaceous materials left on the regenerated or partiallyregenerated catalyst, the weight hourly space velocity (WHSV), theamount of quench medium used, and other relevant process designcharacteristics.

[0030] A wide range of WHSV, defined as weight of total oxygenatefeedstock per hour per weight of catalyst, for the feedstock willfunction in the present invention. Depending on the reactor type, thedesired conversion level, the feedstock composition, and other reactionparameters, the WHSV generally should be in the range of from about 0.01hr⁻¹ to about 1000 hr⁻¹, preferably in the range of from about 0.1 hr⁻¹to about 500 hr⁻, and most preferably in the range of from about 0.5hr⁻¹ to about 200 hr⁻. Since the catalyst may contain other materialswhich act as inerts, fillers, or binders; the WHSV is calculated only onthe weight basis of oxygenate and molecular sieve part of the catalyst.

[0031] One or more diluents may be fed to the reaction zone with theoxygenates, such that the total feed mixture comprises diluent in arange of from about 1 mol % and about 99 mol %. Diluents which may beemployed in the process include, but are not necessarily limited to,helium, argon, nitrogen, carbon monoxide, carbon dioxide, hydrogen,water, paraffins, other saturated hydrocarbons (such as methane, ethane,propane, and mixtures thereof), aromatic compounds, and mixtures thereofPreferred diluents include, but are not necessarily limited to water andnitrogen.

[0032] Oxygenate conversion should be maintained sufficiently high toavoid the need for commercially unacceptable levels of recycling. 100%oxygenate conversion is preferred for the purpose of avoiding feedstockrecycle completely. However, a reduction in unwanted by-products isobserved frequently when the oxygenate, particularly methanol,conversion level is about 98% or less. Accordingly, there is usuallyfrom about 0.05 mol % to about 50 mol % unreacted oxygenate in theproduct stream along with the oxygenate conversion products comprisingolefins, water, and/or other byproducts. It is preferable to recover asmuch of the unreacted oxygenate as possible for recycle purposes. In anyevent, the oxygenate content in waste water may need to be reduced to anenvironmentally acceptable level before byproduct water can bedischarged.

[0033] Therefore, it is desirable to consider this incomplete oxygenateconversion in the overall heat recovery and heat integration scheme,i.e. optimizing heat recovery and heat integration, when using afractionator to recover unreacted oxygenates. If the oxygenateconversion level is high enough and/or recovery of unreacted oxygenateis not warranted for economic or environmental purposes, then thisinvention calls for utilizing heat directly from the heavy productfraction or any or all of the several fractions into which the heavyproduct fraction may be divided.

[0034] After contacting the oxygenate feed, the catalyst becomes fullyor partially deactivated due to accumulation of carbonaceous deposits(also called coke) on the catalyst surface and/or inside the pores.Fully deactivated catalyst The deactivated catalyst having carbonaceousdeposits is separated from the other oxygenate conversion products.Preferably at least a portion of the deactivated catalyst is separatedand withdrawn from the oxygenate conversion reactor intermittently,semi-continuously, continuously, or in batch. Before the deactivatedcatalyst is recycled back to the oxygenate conversion and used again, asuitable regeneration is carried out on at least a portion of thewithdrawn deactivated catalyst to remove at least a portion of thecarbonaceous deposits, in the range of from about 0.1 wt % to about 99.9wt %, preferably at least about 1.0 wt % of the carbonaceous depositsshould be removed. Complete regeneration——removing 100 wt % of theoriginal carbonaceous deposits on all of the deactivated catalyst—alsomay be carried out, but it is found that complete regeneration has atendency of leading to production of large amounts of undesirablebyproducts such as methane and/or hydrogen.

[0035] Preferably, the regeneration is carried out in the presence of agas comprising oxygen or other oxidants. Air and air diluted withnitrogen, steam, and/or CO₂ are preferred regeneration gases. Thecatalyst regeneration temperature should be in the range of from about250° C. to about 750° C., preferably from about 300° C. to about 700° C.

[0036] Almost any type of reactor will provide some conversions of theoxygenates to olefins. Reactor type includes, but is not necessarilylimited to fluidized bed reactor, riser reactor, moving bed reactor,fixed bed reactor, continuously stirred tank reactor, hybrids andcombinations thereof. Increased heat recovery and improved heatintegration in the present invention can be achieved with most anyreactor types. A preferred reactor system for the present invention is acirculating fluid bed reactor with continuous or semi-continuouscatalyst regeneration, similar to a modem fluid catalytic cracker. Fixedbeds may be used, but are not preferred.

[0037] Because the oxygenate conversion reaction is highly exothermic,the oxygenate conversion reaction product effluent generally has ahigher temperature than the feedstock temperature just before contactingthe catalyst. In one embodiment of the present invention, the feedstockfrom the storage tank at a first temperature and having a first heatcontent is heated through several intermediate stages in heat exchangersto a second desired temperature prior to contacting the oxygenateconversion catalyst. It is preferable to have from one to about threestages of heat exchange to provide streams with successively higher heatcontents. Various streams from the oxygenate conversion process atdifferent temperatures and external sources of heat, such as that fromsteam, may be used as heat exchanger fluids to increase either the heatcontent, the temperature, or both, of the feedstock oxygenate.

[0038] After contacting the oxygenate feedstock with the oxygenateconversion catalyst the oxygenate conversion reaction product effluentcomprising olefin products is quenched directly by contacting a suitablequench medium in a quench tower without first going through a productfractionation step. Alternatively, the product effluent may be used toprovide heat directly to the oxygenate feedstock. The temperature andthe heat content of the product effluent are reduced to intermediatelevels afterwards. The product effluent at this lower temperature andlower heat content is sent to the quench tower for direct quenching.

[0039] The compounds in the effluent stream which are gaseous under thequenching conditions are separated from the quench tower as a lightproduct fraction for olefin product recovery and purification. The lightproduct fraction comprises light olefins, dimethyl ether, methane, CO,CO₂, ethane, propane, and other minor components such as water andunreacted oxygenate feedstock. The compounds in the effluent streamwhich are liquid under quenching conditions, are separated from thequench tower as a heavy product fraction for heat recovery, and possibledivision into several fractions and separation of the quench medium. Theheavy product fraction comprises byproduct water, a portion of theunreacted oxygenate feedstock (except those oxygenates that are gasesunder quenching conditions), a small portion of the oxygenate conversionbyproducts, particularly heavy hydrocarbons (C5+), and usually the bulkof the quench medium.

[0040] Preferably, a quench medium is selected from a composition whichremains substantially as a liquid under the quenching conditions, thusminimizing the amount of the quench medium present in the light gaseousproduct fraction which must undergo more expensive gaseous productprocessing steps to recover commercially acceptable grades of lightolefin products. A preferred quench medium is selected from the groupconsisting of water and streams that are substantially water. Morepreferably, the quench medium is a stream which is substantially waterand is selected from the several fractions of the heavy product fractionfrom the quench tower.

[0041] The amount of quench medium circulated in the quench tower at aparticular temperature for product quenching should be not more thanwhat is needed to produce a heavy product fraction exiting the quenchtower having a temperature at least about 5° C. higher than the firsttemperature of the oxygenate feedstock from the storage tank. In anotherembodiment, as already discussed, the oxygenate conversion reactoreffluent stream is used directly as a heat exchanger fluid to provideheat to the oxygenate feedstock before it enters the oxygenateconversion reactor to contact the oxygenate conversion catalyst.

[0042] Preferably, the pressure in the quench tower and the temperatureof the heavy product fraction effluent are maintained at effectivelevels for recovery of the highest quantity and quality of process heat.More preferably, the difference between the heavy product fractioneffluent pressure and the pressure at which the feedstock is vaporizedis below about 345 kPa, more preferably below about 207 kPa. Thetemperature of the heavy product fraction effluent from the quench towerpreferably is maintained at not less than about 30° C. below the bubblepoint of the heavy product fraction effluent. Maintaining a temperaturedifferential between the heavy product fraction effluent and its bubblepoint provides the highest possible bottoms temperature in the quenchtower and the most economically practical recovery of useful heat fromthe heavy product fraction effluent.

[0043] Preferably, the heavy product fraction effluent (heavy productfraction) from the quench tower is pressurized and used for providingheat to other streams. In one embodiment, the heavy product fraction, orany, or all of the several fractions into which the heavy productfraction is divided, or streams from quench medium separations thereof,are used directly as a heat exchanger fluid to increase the heat contentand/or temperature of the oxygenate feedstock at one or more of thestages with successively higher heat contents. Further any of theseveral fractions or streams produced from the quench medium separationsthereof may be used to increase the heat contents of other streamswithin the overall oxygenate conversion reaction and product recoveryprocess The cooled quench medium recovered from such fractions andstreams may be returned back to the quench tower.

[0044] In a preferred embodiment, particularly when the oxygenateconversion is not complete and the quench medium consists essentially ofwater, the heavy product fraction is divided into two fractions, a firstfraction and a second fraction. The relative quantities of the firstfraction and the second fraction depend on the overall amount of heatthat needs to be removed from the product effluent stream in the quenchoperation, and the temperature of the quench medium introduced into thequench tower. The relative quantities are set to optimize equipment costfor heat recovery and utility consumptions. The first fraction is cooledto a desired temperature and sent back to the quench tower as a recycle,i.e. quench water. The utility required to cool the first fraction, e.g.cooling water, may be reduced by using the product effluent stream fromthe oxygenate conversion reactor as a heat exchange fluid to heat theoxygenate feedstock before the feedstock enters the oxygenate conversionreactor and/or before the product effluent stream enters the quenchtower.

[0045] The second fraction of the heavy product fraction effluent issent to a fractionator to separate the quench medium, which consistsessentially of water——a part of it may originate as the recycled portionof the byproduct water from the oxygenate conversion reaction when thefeedstock oxygenate has at least one oxygen——from other compounds, suchas unreacted oxygenates or certain heavier hydrocarbons from theoxygenate conversion reaction, present in the fraction. If other streamshaving compositions similar to or compatible with the second fractionexist within the oxygenate conversion and the associated productrecovery process, such other streams are combined with the secondfraction first and the combined stream is sent to the fractionator.

[0046] Generally it is desirable to fractionate a mixture intocomponents as sharply as possible. In the present invention, it ispreferable for the overhead oxygenate fraction and/or theheavies-containing fraction from the fractionator to have a compositionof water as introduced in the second fraction of the heavy productfraction in the range of from about 15 mol % to about 99.5 mol %,preferably from about 25 mol % to about 90 mol %. An increase in thewater composition of the overhead fraction tends to increase thecondensation temperature, and more heat can be recovered economicallyfrom the overhead fraction of the fractionator to improve heatintegration for the entire process. Preferably, the recovered overheadoxygenate fraction contains at least about 90 mol % of the oxygenatecontained in the second fraction of the heavy fraction. More preferably,the recovered overhead oxygenate fraction contains at least about 99 mol% of the oxygenate contained in the second fraction of the heavyfraction.

[0047] The overhead fraction of the fractionator is condensed in a heatexchanger, i.e. a condenser, against the oxygenate feedstock at one ofthe stages, from one to about three where the oxygenate feedstock isgiven successively higher heat contents. It s preferable for theoverhead fraction of the fractionator to have a pressure at least about69 kPa higher than the pressure of the oxygen feedstock in thecondenser. This pressure differential also increases the condensationtemperature of the overhead fraction, making heat recovery from theoverhead fraction more economical.

[0048] The bottoms fraction of the fractionator consists essentially ofbyproduct water from the oxygenate conversion reaction. Preferably, thisbottoms fraction is pressurized and used to heat the oxygenate feedstockat one of the stages, from one to about three, where the oxygenatefeedstock is given successively higher heat contents prior to enteringthe oxygenate conversion reactor. The fractionator is operated such thatthe temperature of the bottoms fraction is at least about 5° C.,preferably at least about 25° C., higher than the first temperature ofthe oxygenate feed from storage. The operating temperature inside of thefractionator is determined by a number of parameters, including, but notnecessarily limited to the fractionator overhead pressure and theoverall pressure drop inside of the fractionator.

[0049]FIG. 1 shows one embodiment of a process flow diagram according tothe invention to increase heat recovery and to improve heat integration.Liquid oxygenate feed 1, such as methanol, having a first heat content,at a first temperature and a first pressure, is heated by stream 35 inheat exchanger 2. Stream 35 is fractionator bottoms stream 33 fromfractionator 24, which is pressurized by pump 34. The result is a firstheated oxygenate feed stream 3 with a higher heat content than that ofliquid oxygenate feed stream 1. First heated oxygenate feed stream 3then is heated in another heat exchanger 4 by overhead fraction 26 fromfractionator 24 to form a second heated oxygenate feed stream 5 with ahigher heat content than that of stream 3. Heat exchanger 4 is acondensor or a partial condenser for fractionator 24. Second heatedoxygenate feed stream 5 goes through steam pre-heater 6 to form a thirdheated oxygenate feed stream 7 which is further heated by oxygenateconversion product effluent 11 in heat exchanger 8 to form a fourthheated oxygenate feed stream 9 under the effectiveconditions——temperature, pressure, and proportion of liquid andvapor——desired for the conversion of the oxygenate feed. Oxygenateconversion product 11 is the effluent of oxygenate conversion reactor10, after being separated from the deactivated oxygenate conversioncatalyst which has carbonaceous deposits. Alternatively, heat exchanger8 may comprise of a plurality of coils inside of oxygenate conversionreactor 10.

[0050] Fourth heated oxygenate feed stream 9 is fed to oxygenateconversion reactor 10 which contains catalyst suitable for convertingthe oxygenate feed to olefins. Oxygenate conversion reactor 10 may adoptvarious configurations——fixed bed, fluidized bed, riser, moving bed, ora combination thereof, with or without continuous catalyst regeneration.A fixed bed reactor normally is not favored due to the difficulty ofwithdrawing deactivated catalyst for regeneration and returning theregenerated catalyst back to the reactor. The oxygenate feed isconverted to a product comprising light olefins and the catalyst becomesdeactivated or partially deactivated by accumulating carbonaceousdeposits which are formed as byproducts of the oxygenate conversionreaction.

[0051] Oxygenate conversion product effluent 11 flows through heatexchanger 8 and becomes cooled oxygenate conversion product effluentstream 12 which is sent to quench tower 13. Alternately, heat exchanger8 may be eliminated and oxygenate conversion product effluent 11 is sentdirectly to quench tower 13 without intermediate cooling. In quenchtower 13 oxygenate conversion product stream 12 contacts directly with aquench medium consisting essentially of water at an initial temperatureover a series of suitable contacting devices. The amount of the quenchmedium needed in quench tower 13 is dictated by a number of factors,including, but not necessarily limited to the composition of the quenchmedium, the temperature of quench medium recycle introduced to quenchtower 13, and desired temperature differences and pressure differencesbetween various streams. These differences are discussed whereappropriate. The gaseous products are separated as light productfraction stream 14. Heavy product fraction stream 15, which exits fromthe bottom of the quench tower at an exiting temperature, comprises thebulk of byproduct water, a portion of the unreacted oxygenate feedstock(except those oxygenates that are gaseous under the quenchingconditions), a small portion of the oxygenate conversion byproducts,particularly heavy hydrocarbons (C5+), and usually the bulk of thequench medium.

[0052] A preferred quench medium is water, which is for all intents andpurposes indistinguishable from byproduct water. This eliminates theneed for steps to separate the quench medium from byproduct water in theheavy product fraction. In the event that a quench material other thanwater is used and this quench material is substantially in a liquid formunder quenching conditions, heavy product fraction 15, or any, or all ofthe several fraction into which the heavy product fraction is dividedmay be processed to separate the quench medium from byproduct water. Forexample, if the quench medium is a high boiling hydrocarbon such asdiesel fuel or similar streams, it is immiscible with byproduct water.Such a quench medium can be separated by a properly designed weir systemin the bottom of quench tower 13, or in an API separator or othersimilar devices at many different points of the process in the presentinvention. Further, if any heavy hydrocarbons (C5+) are formed in theoxygenate conversion reaction, they also may be removed from byproductwater in stream 15 or other points in the process in substantially thesame manner as or along with the removal of the quench medium. If thequench medium is a relatively light material which is substantiallygaseous under the quenching conditions, and hence being present insubstantial quantities in the light product fraction, such a quenchmedium can be separated in downstream olefin recovery processesencompassing the entire oxygenate conversion and olefin recovery andpurification process.

[0053] Regardless, the exiting pressure of heavy product fraction stream15 should be less than about 345×10³ pascals (345 kPa) below thepressure of liquid oxygenate feed i Preferably, the exiting temperatureof heavy product fraction stream 15 is maintained at not less than about25° C. below the bubble point of byproduct water in stream 15. Apreferred pressure difference between heavy product fraction stream 15(lower pressure) and liquid oxygenate feed 1 (higher pressure) is lessthan 207 kPa.

[0054] Heavy product fraction stream (quench tower bottoms stream) 15may be used to provide heat to the oxygenate feedstock in heatexchangers 2 4, and/or 6 to increase the heat content of the feedstock.The oxygenate feedstock contains successively higher heat contents atthese stages. One or more of these stages also may be eliminated.Preferably, quench tower bottoms stream 15 is divided into to twofractions, recycle fraction 18 and fractionator feed fraction 21.Recycle fraction 18, a quench water recycle stream, is cooled inexchanger 19 and recycled as quenching stream 20 back to quench tower13. Alternatively, recycle fraction 18 or 20 may be split further intoseveral fractions and these fractions may be cooled to differenttemperatures in different heat exchangers. These fractions, or some ofthem, at different temperatures may be introduced into quench tower 13at different points to better integrate heat recovery and minimizeutility consumption. The heat content of fraction 18 may be used toprovide heat to the oxygenate feedstock in the heat exchanger 2, 4,and/or 6, or at different locations of the entire oxygenate conversionand olefin recovery and purification process to provide heat and toincrease heat recovery.

[0055] Fractionator feed fraction 21, optionally mixed with other watercontaining streams 22, is sent to fractionator 24. At least two streams,fractionator overhead stream 26 and fractionator bottoms stream 33, arefractionated from fractionator feed fraction 21. Fractionator overheadstream 26 should contain at least about 15 mol %, preferably at leastabout 25 mol %, of water from the oxygenate conversion reaction.Conjunctively with or alternatively to this composition preference, thetemperature of fractionator overhead stream 26 should be at least about10° C. higher than the boiling temperature of the oxygenate feed underthe conditions of heat exchanger 4.

[0056] Sufficient heat is added to fractionator 24 via reboiler 25,which when coupled with a sufficient number of trays in fractionator 24results in producing fractionator bottoms stream 33 which comprisessubstantially all byproduct water and quench medium introduced withstream 23.

[0057] Preferably, the quench medium is water. When water is used as thequench medium, bottoms stream 33 consists essentially of the bulk ofbyproduct water from the oxygenate conversion reaction and no furthersteps are necessary to separate byproduct water from the quench medium.If the quench medium is a material other than water and has notpreviously been separated from byproduct water prior to introductioninto the quench tower, this quench material may be separated frombyproduct water in bottoms stream 33, or later in the process asdescribed above. Further, if any heavy hydrocarbons (C5+) are formed inthe oxygenate conversion process, they also may be removed frombyproduct water in stream 33, or later in the process in substantiallythe same manner as or along with the removal of the quench medium.

[0058] Fractionator bottoms stream 33, before leaving fractionator 24,is at a temperature which is at least about 5° C., preferably at leastabout 25° C., higher than the first temperature of the oxygenate feedintroduced from storage 1 to heat exchanger 2. The pressure at the topof fractionator 24 should be at least 69 kPa higher than the pressure inheat exchanger 4 to increase heat recovery. Stream 35 is used to heat upliquid oxygenate feedstock 1 in heat exchanger 2. For better heatrecovery, exiting stream 36 from heat exchanger 2 preferably has atemperature equal to or less than about the temperature of stream 21.

[0059] One way to further improve heat integration and to increase heatrecovery is to use fractionator overhead stream 26 as the heat sourcefor heat exchanger 4. The cooled fractionator overhead stream 27 may befractionated further in separator 28 into vapor discharge stream 29 andliquid reflux 30 which is sent back to fractionator 24 after pressureadjustment with pump 31. It is important to maintain cooled fractionatoroverhead stream 27 at a temperature above the boiling point of the firstheated oxygenate feed 3 to provide favorable heat transfer.

[0060] Another embodiment of this invention relates to convertingoxygenates to olefins with high energy and capital efficiency. In thispreferred embodiment, an indirect heat transfer device for transferringheat from at least a portion of an effluent of an oxygenate conversionreactor to the feedstock is used to cause at least a portion of thefeedstock to vaporize. For example, in FIG. 1, the indirect heattransfer device would be heat exchanger 8 and the effluent of theoxygenate conversion reactor is stream 10. In a preferred embodiment,the heat transfer device is a thermosiphon utilizing a disengaging orcirculation drum.

[0061] As defined herein, the boiling range of a feedstock is thedifference in the temperature of that feedstock at its dewpoint and itsbubble point at any one pressure of operation of the feedstock withinthe heat transfer device. If the feedstock is a single componentfeedstock, e.g., a substantially pure methanol, the feedstock has asingle temperature boiling point at any one pressure. Thus it has a 0°C. boiling range. Typically, the feedstock will comprise one or morecomponents having different boiling points, e.g., methanol, diluents,and other hydrocarbon components. Typically, the feedstock will havemore than one component, and have a boiling point range of at least 2°C.

[0062] In a preferred embodiment, the boiling range is not greater thanabout 30° C., 25° C., 20° C., 15° C., 10° C. and most preferably notgreater than about 5 to 6° C. The oxygenate conversion reactor convertsat least a portion of the feedstock into the effluent, and the effluentcomprises the olefins.

[0063] Desirably, the temperature of the reactor effluent is at least300° C., more preferably at least 350° C., and most preferably at least400° C. The temperature of the reactor effluent preferably is belowabout 700° C. to achieve attractive yields in the oxygenate conversionreaction.

[0064] In another embodiment, a portion of the heat in the reactoreffluent emanating from the indirect heat transfer device vaporizing thefeedstock is used to increase the sensible heat of the feedstock priorto vaporization of the feed stock. This heat is particularly useful foradding sensible heat (which is increasing the heat content to a liquidwithout vaporizing it, so the feedstock must be below its bubble point)to the feed stock.

[0065] In another embodiment, the sensible heat is provided in aseparate indirect heat transfer device from that used for vaporization.Preferably, the sensible heat is added to the feedstock in a separateindirect heat transfer device prior to introducing it to thevaporization indirect heat transfer device.

[0066] In another embodiment, the reactor effluent provides heat ofvaporization or increase in sensible heat, to more than one feedstock inseparate indirect heat transfer devices. The indirect heat transferdevices that can be used to transfer heat include, for example, tubularexchangers, fin-type exchangers, condensers, scraped-surface exchangers,agitated vessels and thermosiphon-boilers. A thermosiphon-boiler is adevice wherein natural circulation of the boiling medium is obtained bymaintaining sufficient liquid head to provide for circulation, i.e.,circulation of feedstock through the device occurs by densitydifferences and is not forced by pumps. A separate vessel, desirably adrum in fluid communication with the exchanger, can be used to receivethe partially vaporized feedstock exiting the exchanger, separate theliquid and vapor, and return the liquid to the entrance of theexchanger. Fresh liquid or partially liquid feedstock can also be sentto the seperate vessel, rather than directly to the heat exchanger.

[0067] The thermosiphon-boiler is a particularly useful embodiment forvaporization of feedstock. Tubular exchangers include ashell-and-tube-type heat exchanger, a U-tube heat exchanger, apacked-lantern-ring exchanger, a outside-packed floating-head exchanger,an internal floating-head exchanger, a bent-tube fixed-tube-sheetexchanger, a bayonet-tube exchanger, a spiral-tube exchanger, afalling-film exchanger and Teflon-head exchanger.

[0068] The temperature of cooled reactor effluent is preferably not lessthan 30° C. below the water dewpoint In the method of this invention,determining the boiling range of a feedstock and the water dewpoint of acooled reactor effluent, and many other useful thermodynamic propertiesof the materials utilized, is desireably determined by the SoaveModified Redlich-Kwong (SMRK) equation of state. The SMRK equation isreadily available in SimSci PRO/II, which is a computer program forchemical process simulation.

[0069] In another embodiment, the reactor effluent stream can be split,using one part to vaporize feedstock in one device. Other parts may beused to provide heat to other materials in other devices. In a preferredembodiment, one part of the reactor effluent may be used to vaporizeoxygenate and another part may be used to vaporize diluent (or a mix ofdiluent and oxygenate) in another device, in a parallel approach,feeding both streams to the reactor. In another embodiment, all or partof the reactor effluent may be used to in one device to vaporizediluent, then the reactor effluent from that device is sent to anotherdevice to vaporize oxygenate (or a mix of diluent and oxygenate), or theorder of vaporization may be reversed, in a series approach. The bestapproach will be determined based upon the selection of oxygenate anddiluent employed, and other economic and process criteria, such asdesired energy efficiency and capital return requirements, usingengineering design and economic principles well known to those skilledin the art and not discussed in detail here.

[0070] In other embodiments, the diluent can have a normal boiling pointbetween −20° C. and 130° C.; preferably −6° C. and 100° C.; and mostpreferably 35° C. and 90° C. The diluent is desirably an aliphatic andaromatic hydrocarbons; C₄ to C₈ aliphatic and olefinic hydrocarbons andC₆-C₈ aromatics. Most preferably, the diluent is iso- or normal hexane.

[0071] It is particularly beneficial for the diluent to have arelatively low normal boiling point, i.e., bubble point at oneatmosphere pressure, to allow it to be vaporized at a temperature underthe operating conditions of the vaporizer that is low enough to providehigh temperature differentials with the reactor effluent. It should not,however, be so low that it cannot remain a liquid at the operatingpressures of vaporizer and reactor. Especially useful is when thediluent has a normal boiling point close to that of the oxygenate feed,so that it can be mixed with the oxygenate in large proportion, ifdesired, and still maintain a low boiling range. This may make itpossible to perform as much of the heating of the feedstock, bothsensible heat and vaporization of both oxygenate and diluent, in as fewexchangers as possible.

[0072] The preferred diluents have desirable boiling ranges and normalboiling points. An added attractive feature is they have relatively lowheats of vaporization (compared to oxygenates and water/steam). Thiswould allow the heat from the reactor effluent to be most effectivelyutilized in vaporizing the feedstock sent to the reactor, particularlythat fraction which may be obtained in condensing the water out of thereactor effluent, if so chosen. Iso- and normal hexane are particularlyuseful diluents when the oxygenate is methanol.

[0073] The process can be operated with a feedstock in which 0.1-100% ofthe feedstock is in a liquid state. In other variations, the feedstockcould be at least 10%, 20%, 30%, 40%, 50%, 60%, 70%, 80% and 90% in aliquid state. Preferably, the temperature of the feedstock is not lessthan 30° C. lower than its bubble point at the location, and hence atthe pressure, it is introduced to the heat exchange device in which itwill be vaporized, more preferably not less than 20° C. lower, stillmore preferably not less than 10° C. lower, and most desirably no morethan 6° C. lower.

[0074] The invention will be better understood with reference to thefollowing example, which illustrate, but should not be construed aslimiting the present invention.

EXAMPLE I

[0075] A liquid methanol feed 1 at about 386.1 kPa pressure and 38° C.absorbs heat to increase its heat content in heat exchanger 2 fromstream 35, at 158° C. and 1,276 kPa pressure, from methanol/waterfractionator 24 to form the first heated methanol feed stream 3 at atemperature of about 100° C. and a pressure of351.6 kPa. The firstheated methanol feed stream 3 with 4,722 kJ/mole heat content absorbsheat from the fractionator overhead stream 26 in the heat exchanger 4 toform the second heated methanol feed stream 5 with a heat content of6,521 kJ/mole. Stream 5 is heated further by steam in heat exchanger 6to form the third heated methanol feed stream 7 which has even higherheat content than the third heated methanol feed stream 7- 7,390kJ/mole. The third heated methanol feed stream 7 is heated in heatexchanger 8 to form the fourth heated methanol feed stream 9 with themethanol conversion product effluent 11 from the oxygenate conversionreactor 10. The fourth heated methanol feed stream 9, having a muchhigher heat content of 17,102 kJ/mole, is suitable for contacting acatalyst in the oxygenate conversion reactor 10 to form a deactivatedoxygenate conversion catalyst having carbonaceous deposits and a product11 comprising olefins, particularly light olefins. The oxygenateconversion reactor 10 is a fluidized bed reactor with continuouscatalyst regeneration and recirculation (not shown). The oxygenateconversion product 11 is separated from deactivated oxygenate conversioncatalyst having carbonaceous deposits and used to heat the stream 9 andform a cooled methanol conversion product stream 12. A part of thedeactivated catalyst is withdrawn and removed for regeneration. (notshown). It is preferable to remove at least about 1.0 wt % of thecarbonaceous deposits from the deactivated catalyst during theregeneration. It is also preferred to remove less than about 98.0 wt %of the carbonaceous deposits from the deactivated catalyst duringregeneration. The regenerated catalyst is recycled back into theoxygenate conversion reactor 10 for contacting the oxygenate feed. 99.8wt % of the methanol in stream 9 is converted in the reactor 10, withthe unconverted balance exiting in the stream 11.

[0076] The cooled methanol conversion product stream 12 exiting the heatexchanger 8 is sent to the quench tower 13, contacting directly a quenchmedium consisting essentially of water. The quench tower 13 is equippedwith suitable contacting devices inside. Most hydrocarbon products areseparated as a gaseous product stream 14. Heavier products, water, andunreacted methanol are discharge from the quench tower 13 as the quenchtower bottoms stream 15 at a temperature of about 116° C. and a pressureof about 262 kPa. The quench tower bottoms stream 15 is pressurized bythe pump 16 to form the pressurized quench tower bottoms stream 17 atabout 689.5 kpa. About 83 mol % of the pressurized quench tower bottomsstream 17 forms the recycle fraction 18 and is sent through the coolingexchanger 19 to form the quenching stream 20 at a lower temperature. Thequenching stream 20 is returned to the quench tower 13. The rest of thepressurized quench tower bottoms stream 17, about 17 mol % becomes thefractionator feed fraction 21. The fractionator feed fraction 21 iscombined with another methanol/water stream 22, a small stream recoveredfrom other sources within the overall oxygenate conversion and productrecovery process. The combined stream 23 is sent to the fractionator 24.The fractionator overhead stream 26 containing about 89 mol % water andabout 10.5 mol % of methanol at a temperature of 152° C. and a pressureat 551.6 kPa is sent to the heat exchanger 4. The bottoms fromfractionator 24 is heated with steam in the heat exchanger 25 to producethe fractionator bottoms stream 33 at 158° C. and about 585.4 kPa, whichcontains primarily water with only traces of other components. Thefractionator bottoms stream 33 is pressurized to about 1274.8 kPa andthe resulting stream 35 is used for the heat exchanger 2 to heat theliquid methanol feed 1. After heat exchange, the byproduct warm waterstream 36 has a temperature of 46° C. at a pressure of 861.2 kPa.

[0077] Table 1 shows the product selectivity and the composition ofproduct stream 11 of methanol conversion used for obtaining the resultsshown in Table 2 and Table 3. The feed rates, compositions, pressures,and temperatures of various streams as described in Example I are shownin Table 2. The duties of key exchangers 2, 4, and 25 are tabulated inTable 3. Tables 2 and 3 were compiled using the Simulation Sciences,Inc. PRO/II chemical process simulation program utilizing the ModifiedPanagiotopoulos-Reid modifications to the Soave-edlich-Kwong equation ofstate. TABLE 1 Product Selectivity Hydrocarbon Composition Component (wt%) in Stream 11 (mol %) Hydrogen 0.15 0.73 Carbon Monoxide 0.03 0.01Carbon Dioxide 0.12 0.03 Methane 1.00 0.61 Ethylene 40.90 14.40 Ethane0.83 0.27 Propylene 40.90 9.60 Propane 0.21 0.05 Butenes 8.89 1.50Butanes 0.09 0.02 Pentenes 3.95 0.56 Pentanes 0.04 0.01 Coke 2.89 —Total 100.00 27.84

[0078] TABLE 2 Temper- Heat Stream Rate Methanol Water Pressure atureContent No. (mol/h) (mol %) (mol %) (kPa) (° C.) (kJ/mol)  1 10,000.098.23 1.77 386.1 35.2 582  3 10,000.0 98.23 1.77 351.6 100.1 4.722  510,000.0 98.23 1.77 330.9 98.2 6,521  7 10,000.0 98.23 1.77 317.2 96.97,390  9 10,000.0 98.23 1.77 317.2 96.9 17,102 11 13,918.9 0.14 72.02275.8 407.9 25,424 12 13,918.9 0.14 72.02 262.0 124.9 18,475 14 3,946.70.04 1.78 241.3 37.8 6,841 15 92,909.4 0.18 99.82 262.0 115.6 3,856 219,972.2 0.18 99.82 689.5 115.6 3,859 22 863.8 0.21 99.78 689.5 43.41,394 26 1,118.0 10.54 89.42 551.6 151.9 21,740 27 1,118.0 10.54 89.42517.1 138.8 5,566 29 53.9 36.42 62.75 517.1 138.8 886 33 10,782.2 trace100.00 585.4 157.9 5,302 35 10,782.2 trace 100.00 1,274.8 158.1 5,310 3610,782.2 trace 100.00 861.2 46.1 1,476

[0079] TABLE 3 Exchanger No. Duty (10⁶kJ/h) 2 41.4 4 18.9 6 8.6 8 97.119 191.7 25 38.1

[0080] These results show that in the oxygenate conversion process, theexternal heat needed to bring the oxygenate feedstock to conditionsdesirable for contacting the catalyst, represented in the preferredembodiment by heat exchanger 6, is reduced as a result of increased heatrecovery and improved heat integration of the process.

[0081] Persons of ordinary skill in the art will recognize that manymodifications may be made to the present invention without departingfrom the spirit and scope of the present invention. The embodimentdescribed herein is meant to be illustrative only and should not betaken as limiting the invention, which is defined by the followingclaims.

We claim:
 1. A process for converting an oxygenate to olefins, theprocess comprising: providing a feedstock comprising the oxygenate, andtransferring heat from at least a portion of an effluent of an oxygenateconversion reactor to the feedstock to cause at least a portion of thefeedstock to vaporize and form a vaporized feedstock, wherein thefeedstock has a boiling range of no greater than about 30° C., theoxygenate conversion reactor converts at least a portion of thefeedstock into said effluent, and said effluent comprises said olefins.2. The process of claim 1, wherein the feedstock further comprises adiluent.
 3. The process of claim 1, wherein the feedstock is selectedfrom the group consisting of an impure methanol having a boiling rangeand a mixture of methanol and dimethyl ether.
 4. The process of claim 1,wherein the oxygenate is selected from the group consisting of methanol,dimethyl ether, ethanol, methyl ethyl ether, dimethyl carbonate, methylformate, methyl acetate, diethyl ether, and mixtures thereof.
 5. Theprocess of claim 1, wherein the catalyst is selected from the groupconsisting of a zeolite, a silicoaluminophosphate (SAPO), a substitutedsilicoalumino-phosphate, a substituted aluminophosphate and mixturesthereof.
 6. The process of claim 5, wherein the catalyst is a SAPOselected from the group consisting of SAPO-17, SAPO-18, SAPO-34,SAPO-44, and mixtures thereof.
 7. The process of claim 2, wherein thediluent is a hydrocarbon selected from the group consisting of C4 to C8olefin, C4 to C8 aliphatic, C6-C8 aromatic and mixtures thereof.
 8. Theprocess of claim 2, wherein the diluent is a hydrocarbon selected fromthe group consisting of iso- and normal hexane and mixtures thereof. 9.The process of claim 2, wherein the diluent has a normal boiling pointof about 35° C. to about 90° C.
 10. The process of claim 1, wherein theboiling range is no greater than about 20° C.
 11. The process of claim1, wherein the boiling range is no greater than about 10° C.
 12. Theprocess of claim 1, wherein a temperature of the effluent beforetransferring heat from the effluent to the feedstock is at least about350° C.
 13. The process of claim 1, wherein 0.1-100% of the feedstock isin a liquid state.
 14. The process of claim 1, wherein at least about80% of the feedstock is in a liquid state.
 15. The process of claim 1,wherein at least about 20% of the feedstock is in a liquid state. 16.The process of claim 1, wherein a temperature of the feedstock is notless than 30° C. lower than the bubble point of the feedstock at alocation where the feedstock is introduced to a heat transfer devicethat transfers heat from the effluent to the feedstock.
 17. The processof claim 1, wherein at least a portion of the heat transferred from theeffluent to the feedstock provides sensible heat.
 18. The process ofclaim 1, wherein a temperature of the effluent after transferring heatto the feedstock is not less than 30° C. below water dewpoint.
 19. Theprocess of claim 1, wherein the heat transferred from the effluent tothe feedstock is done in an indirect heat transfer device.
 20. Theprocess of claim 19, wherein the indirect heat transfer device isselected from the group consisting of a shell-and-tube-type exchangerand a thermosiphon-boiler.
 21. A process for converting an oxygenate toolefins, the process comprising: providing a feedstock comprising theoxygenate, transferring heat from at least a portion of an effluent ofan oxygenate conversion reactor to the feedstock to cause at least aportion of the feedstock to vaporize and form a vaporized feedstock, andcontacting the vaporized feedstock at a temperature from about 200 toabout 750° C. and a pressure from 1 kPa to 100 MPa with a catalystcomprising a molecular sieve having a pore diameter smaller than 10Angstroms, wherein the feedstock has a boiling range of no greater thanabout 30° C., the oxygenate conversion reactor converts at least aportion of the feedstock into said effluent, and said effluent comprisessaid olefins.
 22. The process of claim 21, wherein the feedstock furthercomprises a diluent.
 23. The process of claim 21, wherein the feedstockis selected from the group consisting of an impure methanol having aboiling range and a mixture of methanol and dimethyl ether.
 24. Theprocess of claim 21, wherein the oxygenate is selected from the groupconsisting of methanol, dimethyl ether, ethanol, methyl ethyl ether,dimethyl carbonate, methyl formate, methyl acetate, diethyl ether, andmixtures thereof.
 25. The process of claim 21, wherein the catalyst isselected from the group consisting of a zeolite, asilicoaluminophosphate (SAPO), a substituted silicoalumino-phosphate, asubstituted aluminophosphate and mixtures thereof.
 26. The process ofclaim 25, wherein the catalyst is a SAPO selected from the groupconsisting of SAPO-17, SAPO-18, SAPO-34, SAPO-44, and mixtures thereof.27. The process of claim 22, wherein the diluent is a hydrocarbonselected from the group consisting of C4 to C8 olefin, C4 to C8aliphatic, C6-C8 aromatic and mixtures thereof.
 28. The process of claim22, wherein the diluent is a hydrocarbon selected from the groupconsisting of iso- and normal hexane and mixtures thereof.
 29. Theprocess of claim 22, wherein the diluent has a normal boiling point ofabout 35° C. to about 90° C.
 30. The process of claim 21, wherein theboiling range is no greater than about 20° C.
 31. The process of claim21, wherein the boiling range is no greater than about 10° C.
 32. Theprocess of claim 21, wherein a temperature of the effluent beforetransferring heat from the effluent to the feedstock is at least about350° C.
 33. The process of claim 21, wherein 0.1-100% of the feedstockis in a liquid state.
 34. The process of claim 21, wherein at leastabout 80% of the feedstock is in a liquid state.
 35. The process ofclaim 21, wherein at least about 20% of the feedstock is in a liquidstate.
 36. The process of claim 21, wherein a temperature of thefeedstock is not less than 30° C. lower than the bubble point of thefeedstock at a location where the feedstock is introduced to a heattransfer device that transfers heat from the effluent to the feedstock.37. The process of claim 21, wherein at least a portion of the heattransferred from the effluent to the feedstock provides sensible heat.38. The process of claim 21, wherein a temperature of the effluent aftertransferring heat to the feedstock is not less than 30° C. below waterdewpoint.
 39. The process of claim 21, wherein the heat transferred fromthe effluent to the feedstock is done in an indirect heat transferdevice.
 40. The process of claim 39, wherein the indirect heat transferdevice is selected from the group consisting of a shell-and-tube-typeexchanger and a thermosiphon-boiler.
 41. A process for converting anoxygenate to olefins, the process comprising: providing a feedstockcomprising the oxygenate and a diluent, transferring heat from at leasta portion of an effluent of an oxygenate conversion reactor to thefeedstock to cause at least a portion of the feedstock to vaporize andform a vaporized feedstock, and separating the diluent from saideffluent, wherein the feedstock has a boiling range of no greater thanabout 30° C., the oxygenate conversion reactor converts at least aportion of the feedstock into said effluent, and said effluent comprisessaid olefins.
 42. The process of claim 41, wherein the feedstock isselected from the group consisting of an impure methanol having aboiling range and a mixture of methanol and dimethyl ether.
 43. Theprocess of claim 41, wherein the oxygenate is selected from the groupconsisting of methanol, dimethyl ether, ethanol, methyl ethyl ether,dimethyl carbonate, methyl formate, methyl acetate, diethyl ether, andmixtures thereof.
 44. The process of claim 41, wherein the catalyst isselected from the group consisting of a zeolite, asilicoaluminophosphate (SAPO), a substituted silicoalumino-phosphate, asubstituted aluminophosphate and mixtures thereof.
 45. The process ofclaim 44, wherein the catalyst is a SAPO selected from the groupconsisting of SAPO-17, SAPO-18, SAPO-34, SAPO-44, and mixtures thereof.46. The process of claim 41, wherein the diluent is a hydrocarbonselected from the group consisting of C4 to C8 olefin, C4 to C8aliphatic, C6-C8 aromatic and mixtures thereof.
 47. The process of claim41, wherein the diluent is a hydrocarbon selected from the groupconsisting of iso- and normal hexane and mixtures thereof.
 48. Theprocess of claim 41, wherein the diluent has a normal boiling point ofabout 35° C. to about 90° C.
 49. The process of claim 41, wherein theboiling range is no greater than about 20° C.
 50. The process of claim41, wherein the boiling range is no greater than about 10° C.
 51. Theprocess of claim 41, wherein a temperature of the effluent beforetransferring heat from the effluent to the feedstock is at least about350° C.
 52. The process of claim 41, wherein a temperature of theeffluent before transferring heat from the effluent to the feedstock isat least about 300° C.
 53. The process of claim 41, wherein 0.1-100% ofthe feedstock is in a liquid state.
 54. The process of claim 41, whereinat least about 80% of the feedstock is in a liquid state.
 55. Theprocess of claim 41, wherein at least about 20% of the feedstock is in aliquid state.
 56. The process of claim 41, wherein a temperature of thefeed stock is not less than 30° C. lower than the bubble point of thefeedstock at a location where the feedstock is introduced to a heattransfer device that transfers heat from the effluent to the feedstock.57. The process of claim 41, wherein at least a portion of the heattransferred from the effluent to the feedstock provides sensible heat.58. The process of claim 41, wherein a temperature of the effluent aftertransferring heat to the feedstock is not less than 30° C. below waterdewpoint.
 59. The process of claim 41, wherein the heat transferred fromthe effluent to the feedstock is done in an indirect heat transferdevice.
 60. The process of claim 59, wherein the indirect heat transferdevice is selected from the group consisting of a shell-and-tube-typeexchanger and a thermosiphon-boiler.
 61. A process for converting anoxygenate to olefins, the process comprising: providing a feedstockcomprising the oxygenate and a diluent, transferring heat from at leasta portion of an effluent of an oxygenate conversion reactor to thefeedstock to cause at least a portion of the feedstock to vaporize andform a vaporized feedstock, contacting the vaporized feedstock at atemperature from about 200 to about 750° C. and a pressure from 1 kPa to100 MPa with a catalyst comprising a molecular sieve having a porediameter smaller than 10 Angstroms, and separating the diluent from saideffluent, wherein the feedstock has a boiling range of no greater thanabout 30° C., the oxygenate conversion reactor converts at least aportion of the feedstock into said effluent, and said effluent comprisessaid olefins.
 62. The process of claim 61, wherein the feedstock isselected from the group consisting of an impure methanol having aboiling range and a mixture of methanol and dimethyl ether.
 63. Theprocess of claim 61, wherein the oxygenate is selected from the groupconsisting of methanol, dimethyl ether, ethanol, methyl ethyl ether,dimethyl carbonate, methyl formate, methyl acetate, diethyl ether, andmixtures thereof.
 64. The process of claim 61, wherein the catalyst isselected from the group consisting of a zeolite, asilicoaluminophosphate (SAPO), a substituted silicoalumino-phosphate, asubstituted aluminophosphate and mixtures thereof.
 65. The process ofclaim 64, wherein the catalyst is a SAPO selected from the groupconsisting of SAPO-17, SAPO-18, SAPO-34, SAPO-44, and mixtures thereof.66. The process of claim 61, wherein the diluent is a hydrocarbonselected from the group consisting of C4 to C8 olefin, C4 to C8aliphatic, C6-C8 aromatic and mixtures thereof.
 67. The process of claim61, wherein the diluent is a hydrocarbon selected from the groupconsisting of iso- and normal hexane and mixtures thereof.
 68. Theprocess of claim 61, wherein the diluent has a normal boiling point ofabout 35° C. to about 90° C.
 69. The process of claim 61, wherein theboiling range is no greater than about 20° C.
 70. The process of claim61, wherein the boiling range is no greater than about 10° C.
 71. Theprocess of claim 61, wherein a temperature of the effluent beforetransferring heat from the effluent to the feedstock is at least about350° C.
 72. The process of claim 61, wherein a temperature of theeffluent before transferring heat from the effluent to the feedstock isat least about 300° C.
 73. The process of claim 61, wherein 0.1-100% ofthe feedstock is in a liquid state.
 74. The process of claim 61, whereinat least about 80% of the feedstock is in a liquid state.
 75. Theprocess of claim 61, wherein at least about 20% of the feedstock is in aliquid state.
 76. The process of claim 61, wherein a temperature of thefeedstock is not less than 30° C. lower than the bubble point of thefeedstock at a location where the feedstock is introduced to a heattransfer device that transfers heat from the effluent to the feedstock.77. The process of claim 61, wherein at least a portion of the heattransferred from the effluent to the feedstock provides sensible heat.78. The process of claim 61, wherein a temperature of the effluent aftertransferring heat to the feedstock is not less than 30° C. below waterdewpoint as determined by the Soave Modified Redlich-Kwong equation ofstate.
 79. The process of claim 61, wherein the heat transferred fromthe effluent to the feedstock is done in an indirect heat transferdevice.
 80. The process of claim 79, wherein the indirect heat transferdevice is selected from the group consisting of a shell-and-tube-typeexchanger and a thermosiphon-boiler.